Catalytic hydrogenation process

ABSTRACT

A liquid phase catalytic hydrogenation process is described in which an organic feedstock, such as an aldehyde containing from 2 to about 20 carbon atoms, is contracted with hydrogen in the presence of a solid hydrogenation catalyst under hydrogenation conditions to produce a hydrogenation product, such as the corresponding alcohol containing from 2 to about 20 carbon atoms, which process comprises passing a feed solution of the organic feedstock in an inert diluent therefor downwardly in co-current with a hydrogen-containing gas through a hydrogenation zone containing a bed of a particulate hydrogenation catalyst whose particles substantially all lie in the range of from about 1.5 mm to about 5 mm, maintaining the bed of catalyst particles under temperature and pressure conditions conducive to hydrogenation, recovering from a bottom part of the bed a liquid phase containing the hydrogenation product, controlling the rate of supply of the feed solution to the bed so as to maintain a superficial liquid velocity of the liquid down the bed in the range of from about 1.5 cm/sec to about 5 cm/sec, and controlling the rate of supply of the hydrogen-containing gas to the bed so as to maintain at the top surface of the bed of catalyst particles a flow of hydrogen-containing gas containing from 1.00 to about 1.15 times the stoichiometric quantity of hydrogen theoretically necessary to convert the organic feedstock completely to the hydrogenation product.

This invention relates to a liquid phase catalytic hydrogenationprocess.

Heterogeneous catalytic hydrogenation processes of various kinds arewidely practised on a commercial scale and are used for hydrogenation ofa wide variety of organic feedstocks. Typically such hydrogenationreactions are conducted at a pressure of from about 1 bar to about 300bar and at a temperature in the range of from about 40° C. to about 380°C. Examples include hydrogenation of aldehydes to alcohols, ofunsaturated hydrocarbons to saturated hydrocarbons, of acetylene-derivedchemicals to saturated materials, of unsaturated fatty acids tosaturated fatty acids, of ketones to secondary alcohols, of esters ofunsaturated fatty acids to esters of partially or fully hydrogenatedfatty acids, of nitriles to primary amines, and of certain sugars topolyhydroxyalcohols. Also worthy of mention is the hydrogenation ofquinones, for example the hydrogenation of 2-ethylanthraquinone as astep in the production of hydrogen peroxide. This cyclohexanol isproduced commercially by catalytic hydrogenation of cyclohexanone, andiso-propanol by catalytic hydrogenation of acetone. An example ofhydrogenation of an unsaturated hydrocarbon is the production ofcyclohexane from benzene. Typical catalysts for such hydrogenationreactions include Group VIII metal catalysts, such as nickel, palladiumand platinum. Production of butane-1,4-diol by hydrogenation ofbut-2-yn-1,4-diol is an example of hydrogenation of an acetylene-derivedchemical. A suitable catalyst for this reaction has been described as agranular nickel-copper-manganese on silica gel. The production ofstearic acid by catalytic hydrogenation of the corresponding unsaturatedacids, linoleic acid and linolenic acid, at a temperature of about 150°C. and at a pressure of about 14.75 bar to about 32 bar and using anickel, cobalt, platinum, palladium, chromium or copper/zinc catalyst,is an example of the hydrogenation of unsaturated fatty acids to yieldsaturated fatty acids. So-called "hardening" of vegetable oils is anexample of hydrogenation of esters of unsaturated fatty acids.Production of beta-phenylethylamine by hydrogenation of benzyl cyanideis an example of hydrogenation of a nitrile. As examples ofhydrogenation of sugars to polyhydroxyalcohols there can be mentionedhydrogenation of ketose and aldose sugars to hexahydroxyalcohols, forexample hydrogenation of D-glucose to sorbitol and of D-mannose tomannitol.

An important route to C₃ and higher alkanols involves hydroformylationof alpha-olefins, such as ethylene, propylene, and butene-1, to yieldthe corresponding aldehyde having one more carbon atom than the startingolefin. Thus ethylene yields propionaldehyde and propylene yields amixture of n- and iso-butyraldehyde (with the n-isomer usuallypredominating). These aldehydes yield the corresponding alkanols, e.g.n-propanol and n-butanol, upon catalytic hydrogenation. The importantplasticiser alcohol, 2-ethylhexanol, is made by alkali-catalysedcondensation of n-butyraldehyde to yield the unsaturated aldehyde,2-ethyl-hex-2-enal, which is then hydrogenated to yield the desired2-ethylhexanol. Although the preferred catalysts for such aldehydehydrogenation reactions used to be Group VIII metal catalysts, such asnickel, palladium or platinum, the use of a solid catalyst comprising areduced mixture of CuO and ZnO under vapour phase conditions has alsobeen proposed (see EP-A-0008767 and U.S. Pat. No. 2,549,416). Molybdenumsulphide supported on an activated carbon carrier has also beensuggested in GB-A-765972. The hydrogenation of an aldehyde feedcontaining ring-type sulphur compounds using a reduced mixture of oxidesor hydroxides of copper and zinc is described in U.S. Pat. No.4,052,467. Copper chromite has also been used as an aldehydehydrogenation catalyst.

Hydrodesulphurisation is another commercially important hydrogenationreaction. This is the removal complex organic sulphur compounds, such assulphides, disulphides, benzothiophene and the like, from a mixedhydrocarbon feedstock by catalytic reaction with hydrogen to formhydrogen sulphide. In such a process typical operating conditionsinclude use of a temperature of from about 260° C. to about 375° C., ahydrogen pressure of from about 5 bar to about 40 bar and an aluminasupported cobalt-molybdenum or nickel-molybdenum catalyst.

Catalytic hydrogenation is in all the above cases a heterogeneousprocess. It may be operated as a liquid phase process or as a vapourphase process. A review of some of the factors involved in designingheterogeneous gas and vapour phase reaction systems appeared in"Chemical Engineering", July 1955, in an article entitled "Moving Bed--Processes . . . New Applications", at pages 198 to 206 (see inparticular pages 204 and 205 thereof).

There have been various prior proposals to operate hydrogenationprocesses in several catalytic stages connected in series. For example,a vapour phase aldehyde hydrogenation process is described in U.S. Pat.No. 4,451,677 which involves use of a plurality of adiabaticallyoperated catalytic hydrogenation stages connected in series.

DE-B-1115232 describes a process for the production of alcohols with 2to 6 carbon atoms by hydrogenation in the liquid phase over a nickel orcobalt catalyst of a feed mixture comprising the corresponding aldehydediluted with from 50 to 300 volume % of product alcohol, using twohydrogenation stages connected in series. Reaction conditions includeuse of a temperature of 130° C. to 220° C. and a pressure of less than50 bar, whilst the aldehyde feed rate corresponds to a space velocity offrom 0.3 to 2.5 hr⁻¹, preferably 0.75 to 1.1 hr⁻¹. An excess of hydrogenis recirculated from the exit end of the second hydrogenation stage tothe inlet end of the first hydrogenation stage.

GB-A-784359 is concerned with preferential hydrogenation of aldehydes ina mixture of aldehydes and olefins, water being added to inhibit olefinhydrogenation. Multi-bed co-current hydrogenation is used, withinjection of hydrogen between beds. Hydrogen recycle is envisaged.

GB-A-1175709 describes an apparatus for production of cyclohexane bycatalytic hydrogenation of benzene. Excess hydrogen is recycled.

Use of 2-ethylhexanol as solvent to control the temperature duringhydrogenation of a mixture of 2-ethylhexanal and iso-butyraldehyde issuggested in BR-A -PI800154 (Chem. Abs., 96 (1982) 51807h).

CA-A-926847 discloses in Example 2 a process in which a solution of2-ethylanthraquinone is passed through a tubular reactor in co-currentwith hydrogen. U.S. Pat. No. 3,009,782 describes a similar process inwhich the working solution is passed through a fixed bed of thehydrogenation catalyst at a rate of between 20 and 200 liters per minuteper square foot of catalyst bed cross-section (215.3 and 2152.8 litersper minute per square meter of catalyst bed). A further modification ofthis process is outlined in U.S. Pat. No. 3,755,552 which recommendshydrogenation in a hydrogenator shell containing a plurality ofsubstantially vertically oriented, laterally positioned cylinders filledwith catalyst wherein the ratio of the diameter of a cylinder to thediameter of the catalyst particle is at least 15:1.

In conventional liquid phase multi-stage hydrogenation processes thehydrogen-containing gas and the material to be hydrogenated are fedthrough the plant in co-current or in counter-current fashion. In orderto achieve good economy of hydrogen usage it is usual to recycle gaswithin the plant. Hence in designing the plant account must be taken ofthe circulating inert gases (e.g. N₂, Ar, CH₄ and the like) which areinevitably present in the circulating gas of a commercial plant.Moreover, it is recognised in the art that hydrogen is relatively poorlysoluble in organic liquids and so one of the rate limiting steps in aliquid phase hydrogenation process may be the dissolution of hydrogen inthe organic phase and its subsequent migration through the liquid phaseto the catalyst surface. For this reason the use of high partialpressures of hydrogen is often recommended, although often a balance hasto be struck by the plant designer between additional process efficiencyand the additional capital and running costs associated with use of highpressures. An extra factor to be considered is the additional cost ofusing recirculating gas streams at high pressure which containsignificant levels of inert gases as well as hydrogen. Hence the plantdesigner may have to sacrifice efficiency of hydrogen utilisation inorder to avoid the waste of energy involved in recycling inert gases athigh pressures in excess of about 50 bar.

The term trickle bed reactor is often used to describe a reactor inwhich a liquid phase and a gas phase flow concurrently downward througha fixed bed of catalyst particles while reaction takes place. Atsufficiently low liquid and gas flow rates the liquid trickles over thepacking in essentially a laminar film or in rivulets, and the gas flowscontinuously through the voids in the bed. This is sometimes termed thegas continuous region or homogeneous flow and is the type encounteredusually in laboratory and pilot scale operations. As gas and/or liquidflow rates are increased there is encountered behaviour described asrippling, slugging or pulsing flow. Such behaviour may be characteristicof the higher operating rates encountered in commercial petroleumprocessing. At high liquid rates and sufficiently low gas rates, theliquid phase becomes continuous and the gas passes in the form ofbubbles; this is sometimes termed dispersed bubble flow and ischaracteristic of some chemical processing in which liquid flow ratesare comparable to the highest encountered in petroleum processing, butwhere gas/liquid ratios are much less Flow patterns and the transitionsfrom one form to another as a function of gas and liquid flow rates havebeen described by several authors.

A useful general review of trickle bed reactors and other multiphasereactors can be found under the heading "Reactor Technology" in"Kirk-Othmer Encyclopedia of Chemical Technology", Third Edition, Volume19, at pages 880 to 914. This states at page 892:

"Trickle-bed reactors have complicated and as yet poorly defined fluiddynamic characteristics. Contacting between the catalyst and thedispersed liquid film and the film's resistance to gas transport intothe catalyst, particularly with vapor generation within the catalyst, isnot a simple function of liquid and gas velocities Maximum contactingefficiency is attainable with high liquid mass velocities, i.e. 1-5kg/(m² ·s) or higher in all sized units however, 3-8 kg/(m² ·s) is amore preferable range of liquid mass velocities."

Assuming a specific gravity for an organic liquid of approximately 0.8,these liquid velocities indicate that maximum contacting efficiency isattainable at a superficial liquid velocity of 0.24 to 1.0 cm/sec (i.e.3-8 kg/(m² ·s)).

Further reviews of the operation of trickle bed reactors have appearedas follows:

1. "Trickle-bed reactors" by Charles N. Satterfield, AIChE Journal, Vol.21, No. 2 (March 1975), pages 209 to 228;

2. "Chemical Reactor Design for Process Plants" by H. F. Rase (1977),pages 698 to 711;

3. "Multiphase Catalytic Packed-Bed Reactors" by Hanns P. Hofmann, CatalRev.-Sci. Eng., 17(1), pages 71 to 117 (1978);

4. "Encyclopedia of Fluid Mechanics" (1986), Chapter 32 by Milorad P.Dudukovic and Patrick L. Mills, pages 969 to 1017, published by GulfPublishing Company, P. O. Box 2608, Houston, Tex. 77001;

5. "Trickle-Bed Reactors", by Mordechay Herskowitz and J. M. Smith,AIChE Journal, Vol. 29, No. 1 (January 1983) pages 1 to 18;

6. "Hydroprocessing conditions affect catalyst shape selection" by B. H.Cooper, B. B. L. Donnis, and B. Moyse, Technology, Dec. 8, 1986, Oil &Gas Journal, pages 39 to 44;

7. "Gas-Liquid-Solid Reaction Engineering" by Y. T. Shah and D. Smith,IChemE Symposium Series 87 (ISCRE 8);

8. "Trickle-Bed Reactors: Dynamic Tracer Tests, Reaction Studies, andModeling of Reactor Performance" by A. A. El-Hisnawi, M. P. Dudukovicand P. L. Mills, ACS Symposium Series 196, Chemical Reaction Engineering(1982), pages 421 to 440;

9. "Hydrodynamics and interfacial areas in downward cocurrent gas-liquidflow through fixed beds. Influence of the nature of the liquid" by B. I.Morsi, N. Midoux, A. Laurent, and J.-C. Charpentier, InternationalChemical Engineering, Vol. 22, No. 1, pages 142 to 151 (January 1982);

10. "Packing wetting in trickle bed reactors: influence of the gas flowrate" by S. Sicardi, G. Baldi, V. Specchia, I. Mazzarino, and A.Gianetto, Chemical Engineering Science, Vol. 36, pages 226 to 227(1981);

11. "Influence of gas velocity and packing geometry on pulsing inceptionin trickle-bed reactors" by S. Sicardi and H. Hofmann, The ChemicalEngineering Journal, 20 (1980), pages 251 to 253;

12. "Some comments on models for evaluation of catalyst effectivenessfactors in trickle-bed reactors" by P. L. Mills, H. F. Erk, J. Evans,and M. P. Dudukovic, Chemical Engineering Science, (1981), Vol. 36 (5),pages 947 to 950;

13. "Effectiveness Factors and Mass Transfer in Trickle-Bed Reactors" byMordechay Herskowitz, R. G. Carbonell and J. M. Smith, AIChE JournalVol. 25, No. 2 (March 1979) pages 272 to 283;

14. "Flow Regime Transition in Trickle-Bed Reactors" by S. Sicardi, H.Gerhard and H. Hoffmann, The Chemical Engineering Journal, 18 (1979),pages 173 to 182;

15. "Catalyst Effectiveness Factor in Trickle-Bed Reactors" by M. P.Dudukovic and P. L. Mills, Chemical Reaction Engineering--Houston, ACSSymposium Series 65 (1978), pages 387 to 399;

16. "Hydrodynamics and Solid-Liquid Contacting Effectiveness inTrickle-Bed Reactors" by A. Gianetto, G. Baldi, V. Specchia, and S.Sicardi, AIChE Journal, Vol. 24, No. 6, (November 1978), pages 1087 to1104;

17. "Analysis of Three-Phase Packed-Bed Reactors" by S. Goto and J. M.Smith, AIChE Journal, Vol. 24, No. 2, pages 295 to 302;

18. "Performance of Slurry and Trickle-Bed Reactors Application toSulfur Dioxide Removal", by S. Goto and J. M. Smith, AIChE Journal, Vol.24, No. 2, March 1978 pages 286 to 293;

19. "Two-Phase Downflow Through Catalyst Beds: Part 1. Flow Maps" by E.Talmor, AIChE Journal, Vol. 23, No. 6, November 1977, pages 868 to 878;

20. "Pressure Drop and Liquid Holdup for Two Phase Concurrent Flow inPacked Beds" by V. Specchia and G. Baldi, Chemical Engineering Science,Vol. 32, (1977) pages 515 to 523;

21. "Trickle-Bed Reactor Performance: Part 1. Holdup and Mass TransferEffects" by S. Goto and J. M. Smith, AIChE Journal, Vol. 21, No. 4, July1975, pages 706 to 713;

22. "Effect of Holdup Incomplete Catalyst Wetting and Backmixing duringHydroprocessing in Trickle Bed Reactors" by J. A. Paraskos, J. A. Frayerand Y. T. Shah, Ind. Eng. Chem., Process Des. Dev., Vol. 14, No. 3,(1975) pages 315 to 322;

23. "Wetting of Catalyst Particles under Trickle Flow Conditions" by J-BWijffels, J. Verloop and F. J. Zuiderweg, Chemical ReactionEngineering-II, Advances in Chemistry Series, Vol. 133, 1974, pages 151to 163;

24. "The Role of Liquid Holdup and Effective Wetting in the Performanceof Trickle-Bed Reactors" by D. E. Mears, Chemical ReactionEngineering-II, Advances in Chemistry Series, Vol. 133, 1974 pages 218to 227;

25. "Scale Up of Pilot Plant Data for Catalytic Hydroprocessing" by H.C. Henry and J. B. Gilbert, Ind. Eng, Chem. Process Des. Develop., Vol.12, No. 3, 1973, pages 328 to 334;

26. "Direct Solid-Catalyzed Reaction of a Vapor in an ApparentlyCompletely Wetted Trickle Bed Reactor" by C. N. Satterfield and F. Ozel,AIChE Journal, Vol. 19, No. 6, November 1973, pages 1259 to 1261;

27. "Pressure Loss and Liquid Holdup in Packed Bed Reactor withCocurrent Gas-Liquid Down Flow" by Y. Sato, T. Hirose, F. Takahashi, andM. Toda, Journal of Chemical Engineering of Japan, Vol. 6, No. 2, 1973,pages 147 to 152;

28. "Partial Wetting in trickle bed reactors--the reduction ofcrotonaldehyde over a palladium catalyst", by W. Sedriks and C. N.Kenney, Chemical Engineering Science, Vol. 28, 1973, pages 559 to 568;

29. "Handling kinetics from trickle-phase reactors" by A. Bondi, Chem.Tech., March 1971, pages 185 to 188;

30. "Kinetics of Hydrodesulfurization" by C. G. Frye and J. F. Mosby,Chemical Engineering Progress, Vol. 63, No. 9, September 1967, pages 66to 70; and

31. "Performance of Trickle Bed Reactors" by L. D. Ross, ChemicalEngineering Progress, Vol. 61, No. 10, October 1965, pages 77 to 82.

The present invention seeks to provide an improved liquid phasehydrogenation process in which essentially 100% hydrogenation of thealdehyde or other organic feedstock to the desired hydrogenation productcan be achieved, with minimisation of formation of by-products.

It further seeks to provide a liquid phase hydrogenation process inwhich the use of gas recycle compressors is obviated. Additionally itseeks to provide a process for liquid phase hydrogenation of a widevariety of organic feedstocks which can be operated with excellenteconomy of hydrogen usage without the need for recycle ofhydrogen-containing gases.

According to the present invention there is provided a liquid phasecatalytic hydrogenation process in which an organic feedstock iscontacted with hydrogen in the presence of a solid hydrogenationcatalyst under hydrogenation conditions to produce a hydrogenationproduct, which process comprises passing a feed solution of the organicfeedstock in an inert diluent therefor downwardly in co-current with ahydrogen-containing gas through a hydrogenation zone containing a bed ofa particulate hydrogenation catalyst whose particles substantially alllie in the range of from about 0.5 mm to about 5 mm, maintaining the bedof catalyst particles under temperature and pressure conditionsconducive to hydrogenation, recovering from a bottom part of the bed aliquid phase containing the hydrogenation product, controlling the rateof supply of the feed solution to the bed so as to maintain asuperficial liquid velocity of the liquid down the bed in the range offrom about 1.5 cm/sec to about 5 cm/sec, and controlling the rate ofsupply of the hydrogen-containing gas to the bed so as to maintain atthe top surface of the bed of catalyst particles a flow ofhydrogen-containing gas containing from 1.00 to about 1.15 times thestoichiometric quantity of hydrogen theoretically necessary to convertthe organic feedstock completely to the hydrogenation product.

Preferably the catalyst particle size range is from about 0.5 mm toabout 3 mm.

In view of the teaching in the art that, in operation of trickle bedreactors, the maximum gas-liquid contacting efficiency is attainable ata superficial liquid velocity of no more than about 1.0 cm/sec, it ismost surprising to find that, in hydrogenation reactions such as thehydrogenation of an aldehyde to an alcohol, an approximatelystoichiometric quantity of hydrogen, or at most only a minor excess ofhydrogen, can be used to achieve near quantitative hydrogenation in asingle passage over a bed of catalyst of the appropriate depth when thecatalyst particle size range is from about 0.5 mm to about 5 mm and ahigh liquid superficial velocity down the bed, i.e. from about 1.5cm/sec to about 5 cm/sec, is used. Thus, even though the gas near theexit end of the bed may be almost entirely depleted of hydrogen,efficient conversion of unsaturated organic compound (e.g. aldehyde) orother organic feedstock to hydrogenatipn product (e.g. alcohol) can beachieved without having to have recourse to high pressures in excess ofabout 50 bar. Hence the use of a large excess of hydrogen is notnecessary as we have shown, in the course of our experimentation, thatthe influence of hydrogen partial pressure on the rate of hydrogenationis of minor significance. Moreover in our work on hydrogenation ofaldehydes we have found that, under the unconventional flow conditionsused in the process of the invention, high average rates of reaction arepossible, approaching in suitable cases about 5 gm. moles of aldehydehydrogenated per liter of catalyst per hour and at the same timeachieving substantial conversion (i.e. 95% of more) of the aldehyde feedto the alcohol product.

The process of the invention is not specific to any particularhydrogenation reaction or to any particular catalyst composition.However, in general the hydrogenation conditions used in thehydrogenation zone include use of a pressure of from about 1 bar toabout 300 bar, often from about 1 bar to about 100 bar, and of atemperature of from about 40° C. to about 350° C., often from about 90°C. to about 220° C.

In operating the process of the invention a pressure drop is set upacross the catalyst bed, typically of at least about 0.1 kg/cm² permeter of catalyst bed depth. Care must accordingly be taken, indesigning a plant to operate according to the invention, that it isensured that at the bottom of the catalyst bed the crushing strength ofthe catalyst is not equalled or exceeded. If there is any risk of thisoccurring, then it is necessary to utilise two or more catalyst beds ofappropriate depth in place of a single large catalyst bed; in this casegas and liquid must be uniformly distributed into each bed.

The selection of catalyst particle size and of the superficial liquidvelocity are features which are crucial to the process of the invention.These features ensure that the catalyst surface is completely wetted,that a large catalyst superficial surface area is presented for reactionof the unsaturated organic compound or other organic feedstock withhydrogen, that good liquid-gas contact is effected as the gas bubblesentrained in the liquid pass through the irregular channels in the bedin co-current downflow through the bed, that dissolution of hydrogeninto the downflowing liquid is thereby facilitated, and that good masstransfer of the dissolved hydrogen and unsaturated organic compound orother organic feedstock to the catalyst surface is also achieved by therelatively rapid flow of the liquid through the complex network ofinterconnecting passages present in the catalyst bed. In the case ofspherical catalyst particles the actual velocity of the liquid over thecatalyst surface can be up to about 3 times the superficial velocity ofthe gas plus liquid. Another important factor is the concentration ofthe unsaturated organic compound or other organic feedstock in theliquid phase. As hydrogenation is usually an exothermic reaction, theuse of an appropriately dilute solution helps to limit the temperaturerise, particularly when the hydrogenation zone is operated underadiabatic conditions. By selection of an appropriate concentration ofunsaturated organic compound or other organic feedstock in the feedsolution it is possible to optimise hydrogenation conditions at thecatalyst surface so that neither the unsaturated organic compound orother organic feedstock nor any hydrogenation product thereof "blinds"the catalyst to hydrogen. Such "blinding" of the catalyst will occur, itis postulated, if one or more of the species present, whether theunsaturated organic compound or other organic feedstock or somehydrogenation product thereof, is strongly absorbed or adsorbed on thecatalyst surface and thereby prevents approach of hydrogen molecules tothe active catalytic sites.

The process of the invention can be applied, for example to thehydrogenation of unsaturated hydrocarbons to saturated hydrocarbons.Typical of such a reaction is the production of cyclohexane frombenzene. This hydrogenation can be carried out according to theinvention using a nickel, palladium or platinum catalyst in thehydrogenation zone and a temperature of from about 100° C. to about 200°C. and a pressure of from about 5 bar to about 30 bar. This reaction isexothermic. The use of relatively high temperatures is normallyrecommended so as to maximise the rate of conversion of benzene tocyclohexane, but isomerisation of cyclohexane to methyl cyclopentane,which is extremely difficult to separate from cyclohexane, can occur inthe aforementioned conventional procedures, especially at suchrelatively high temperatures.

Production of secondary alcohols by reduction of ketones is anotherappropriate hydrogenation reaction to which the invention can be appliedExamples of such reactions include production of iso-propanol fromacetone and of cyclohexanol from cyclohexanone.

Another example of a hydrogenation reaction to which the presentinvention can be applied is the production of butane-1,4-diol byhydrogenation of but-2-yn-1,4-diol This can be carried out using acatalyst which is a granular nickel-copper-manganese on silica gel at apressure of from about 200 bar to about 300 bar in the hydrogenationzone. A typical inlet temperature to the hydrogenation zone is about 40°C., when the catalyst is freshly reduced

A further example of a hydrogenation reaction to which the process ofthe invention can be applied is the production of stearic acid byhydrogenation of linoleic acid, of linolenic acid, or of a mixturethereof This can be carried out using a nickel, cobalt, platinum,palladium, chromium or zinc catalyst at a pressure of from about 10 barto about 40 bar and an inlet temperature to the hydrogenation zone ofabout 150° C.

Other examples of hydrogenation processes to which the invention can beapplied include "hardening" of vegetable oils, hydrodesulphurization,hydrogenation of nitriles to amines, and hydrogenation of sugars, (forexample, hydrogenation of aldoses, such as D-glucose or D-mannose, tothe corresponding hexahydroxyalcohols, such as sorbitol and mannitol).

A particularly preferred type of hydrogenation reaction is theproduction of alcohols from aldehydes. Such aldehydes generally containfrom 2 to about 20 carbon atoms and may in the case of those aldehydescontaining 3 or more carbon atoms include one or more unsaturatedcarbon-carbon bonds besides the unsaturated --CHO group. Thus as usedherein the term "aldehyde" includes both saturated and unsaturatedaldehydes, that is to say aldehydes wherein the only hydrogenatablegroup is the aldehyde group, --CHO, itself (such as alkanals) andaldehydes which contain further hydrogenatable groups such as olefinicgroups, >C=C<, in addition to the aldehyde group, --CHO (such asalkenals). Typical aldehydes include n- and iso-butyraldehydes,n-pentanal, 2-methylbutanal, 2-ethylhex-2-enal, 2-ethylhexanal,4-t-butoxybutyraldehyde, C₁₀ --"OXO"--aldehydes (e.g.2-propylhept-2-enal), undecanal, dodecanal, tridecanal, crotonaldehydeand furfural, as well as mixtures of two or more thereof. Aldehydes andmixtures of aldehydes can be produced by hydroformylation of an olefinor mixed olefins in the presence of a cobalt catalyst or a rhodiumcomplex catalyst, according to the equation: R·CH═CH₂ +H₂ +CO→R·CH₂ ·CH₂·CHO+R·CH(CHO)·CH₃ ; where R is a hydrogen atom or an alkyl radical. Theratio of the n-aldehyde to the iso-aldehyde in the product depends to acertain extent on the selected hydroformylation conditions and upon thenature of the hydroformylation catalyst used. Although cobalt catalystswere formerly used, more recently the use of rhodium complex catalystshas been preferred since these offer the advantages of lower operatingpressure, ease of product recovery, and high n-iso-aldehyde molarratios. Typical operating conditions for such rhodium complexhydroformylation catalysts can be found in U.S. Pat. No. 3,527,809, U.S.Pat. No. 4,148,830, EP-A-0096986, EP-A-0096987, and EP-A-0096988. Insuch hydroformylation processes the aldehyde or aldehyde products can berecovered in admixture with unreacted olefin and its hydrogenationproduct, i.e. the corresponding paraffin. Such crude reaction productscan be used as starting material in the process of the invention.Further aldehydes can be obtained by condensation reactions; forexample, 2-ethylhex-2-enal can be made by condensation of 2 moles ofn-butyraldehyde and 2-propylhept-2-enal by condensation of 2 moles ofn-valeraldehyde. Examples of aldehyde hydrogenation reactions are theproduction of n-butanol from n-butyraldehyde, of 2-ethylhexanol from2-ethylhex-2-enal, or 2-propylheptanol from 2-propylhept-2-enal, ofundecanol from undecanal, and of 4-t-butoxybutanol from4-t-butoxybutyraldehyde. The invention is used to special advantage forhydrogenation of aldehydes containing from about 7 to about 17 carbonatoms to the corresponding alkanols. In such aldehyde hydrogenationreactions there can be used any of the conventionally used supportedmetal catalysts, such as Ni, Pd or Pt supported on a variety of supportssuch as granular carbon, silica, silica-alumina, zirconia, siliconcarbide or the like, or copper chromite.

Other aldehyde hydrogenation catalysts include cobalt compounds; nickelcompounds which may contain small amounts of chromium or anotherpromoter; mixtures of copper and nickel and/or chromium; and other GroupVIII metal catalysts, such as Pt, Pd, Rh and mixtures thereof, onsupports, such as carbon, silica, alumina and silica-alumina. The nickelcompounds are generally deposited on support materials such as aluminaor kieselguhr.

In all cases the catalyst particles substantially all have a particlesize in the range of from about 0.5 mm to about 5 mm, preferably in therange of from about 0.5 mm to about 3 mm, as measured by a conventionalsieve analysis techhique. By the term "substantially all" we mean thatnot more than about 5%, and preferably not more than about 0.5%, ofparticles are less than about 0.5 mm in size, and that not more thanabout 5%, and preferably not more than about 1%, of particles are largerthan 5 mm (or 3 mm) in size. The catalyst particles may be of anydesired shape, such as cylindrical, but are conveniently approximatelyspherical granules. However the use of pelleted catalysts and ofcatalyst particles of more complex shape is not ruled out. In the caseof spherical or granular catalyst particles the particle size isessentially equivalent to particle diameter, whereas in the case ofcylindrical catalyst particles or particles of more complex shape thesize range refers to the shortest particle dimension, e.g. diameter inthe case of a cylinder or extrudate. Particularly preferred catalystsare those with a particle size range of from about 1 mm to about 2 mm.

The hydrogenation zone may include two or more beds of catalyst.Conveniently, however, the hydrogenation zone comprises a singlecatalyst bed. The depth of the catalyst bed or beds should be sufficientto ensure that the desired degree of conversion (e.g. about 75% to about99% or higher, for example about 99.5% or more) can be effected inpassage through the bed under the selected reaction conditions.

The hydrogen-containing gas supplied to the hydrogenation zonepreferably contains a major amount of hydrogen and at most a minoramount of one or more inert gases, such as nitrogen, methane, other lowmolecular weight hydrocarbons, such as ethane, propane, n-butane andiso-butane, carbon oxides, neoh, argon or the like. Preferredhydrogen-containing gases are accordingly gases containing at leastabout 50 mole % up to about 95 mole % or more (e.g. about 99 mole %), ofH₂ with the balance comprising one or more of N₂, CO, CO₂, Ar, Ne, CH₄and other low molecular weight saturated hydrocarbons. In some cases,for example when using nickel catalysts, the presence of CO and CO₂cannot be tolerated and the total carbon oxides concentration shouldnot, in this case, be more than about 5 to 10 ppm by volume. Suchhydrogen-containing gases can be obtained in conventional manner fromsynthesis gas and other usual sources of hydrogen-containing gases,followed, if necessary, by appropriate pretreatment to removeimpurities, such as sulphurous impurities (e.g. H₂ S, COS, CH₃ SH, CH₃SCH₃, and CH₃ SSCH₃) and halogen-containing impurities (e.g. HCl and CH₃Cl) which would exert a deleterious influence on catalytic activity,i.e. catalyst inhibition, poisoning or deactivation, as well as by theremoval of the carbon oxides. Preparation of suitablehydrogen-containing gases will accordingly be effected according tousual production techniques and forms no part of the present invention.Thus the hydrogen-containing gas supplied to the hydrogenation zone maybe, for example, a 94 mole % hydrogen stream produced by steam reformingof natural gas followed by the water gas shift reaction:

    CO+H.sub.2 O⃡CO.sub.2 +H.sub.2,

then by CO₂ removal to give a gas containing about 1 mole % to about 2mole % carbon oxides, and finally by methanation to give a gascontaining only a few ppm by volume of carbon oxides. Substantially purehydrogen from an electrolysis plant may be used, as can also purifiedhydrogen streams obtained by the pressure swing adsorption treatment ofhydrogen admixed with CO, CO₂ and light hydrocarbon gases, in each casewith excellent results. For a discussion of production of hydrogenstreams by pressure swing adsorption reference may be made to a paperentitled "Hydrogen Purification by Pressure Swing Adsorption" by H. A.Stewart and J. L. Heck, prepared for Symposium on Adsorption--Part III,64th National Meeting of the American Institute of Chemical Engineers,New Orleans, La., U.S.A., March 16-20, 1969.

The rate of supply of the feed solution to the catalyst bed correspondsto a superficial liquid velocity down the bed of from about 1.5 cm/secto about 5 cm/sec, for example from about 1.5 cm/sec to about 3 cm/sec.

The feed solution supplied to the hydrogenation zone contains theunsaturated organic compound or other organic feedstock dissolved in acompatible diluent therefor. The purpose of the diluent is to act as aheat sink, to limit the temperature rise within the hydrogenation zoneto an acceptable limit, and also to provide at the same time anappropriate volumetric flow into the catalyst bed, such that therequired liquid superficial velocity is achieved along with the desiredproduct conversion and temperature rise. The concentration of organicfeedstock in the feed solution is accordingly preferably selected independence on the expected acceptable temperature rise across thehydrogenation zone; such temperature rise should not be so great as tocause more than a minor amount of vaporisation of liquid in thehydrogenation zone or to cause thermal damage to the catalyst, to anyreactant present or to the hydrogenation product.

In a typical process the feed solution supplied to the hydrogenationzone contains at least about 1 mole % of an unsaturated organic compoundup to about 50 mole %, more preferably in the range of from about 5 mole% up to about 33 mole %, the balance being diluent or diluents.

In a typical hydrodesulphurisation process the organic feedstockcomprises one or more organic sulphurous compounds present in ahydrocarbon diluent. The concentration of such sulphurous compounds(expressed as sulphur content) may range from a few ppm, e.g. about 5ppm up to about 10% by weight.

The diluent can be any convenient inert liquid or mixture of liquidsthat is compatible with the unsaturated organic compound or otherorganic feedstock and the catalyst, with any intermediate product orby-product, and with the desired hydrogenation product. In many casesthe hydrogenation product itself can be used as the compatible diluentor as a part of the compatible diluent. Hence, when hydrogenating analdehyde for example, the diluent can be the product alcohol obtainedupon hydrogenation of the aldehyde. In this case the process of theinvention includes the further step of recycling a part of the liquidhydrogenation product for admixture with make up unsaturated organiccompound or other organic feedstock to form the feed solution to thehydrogenation zone. Alternatively aldehyde condensation product, such asthe dimers, trimers and high condensation products of the type disclosedin GB-A-1338237, can be used as diluent. If the unsaturated organiccompound or other organic feedstock used as starting material is a solidor if the hydrogenation product or an intermediate product is a solid,then an inert solvent will usually be used. Similarly, use of a solventmay be desirable in cases in which by-product formation is a problem.For example, hydrazobenzene is a potential by-product of thehydrogenation of nitrobenzene to yield aniline; in such a case it isdesirable to dissolve the unsaturated organic compound, such asnitrobenzene, in a solvent, such as ethanol, in order to limit formationof an undesirable by-product, such as hydrazobenzene. In this case it isalso highly advantageous to include a minor amount of ammonia in theethanol solvent as ammonia further limits the formation of by-productssuch as azobenzene, azoxybenzene or hydroazobenzene.

Because a stoichiometric or near stoichiometric quantity of hydrogen isused in the process of the invention and there is at most only a smallexcess of hydrogen used, the liquid phase hydrogenation of evenrelatively volatile unsaturated organic compounds to similarly volatileproducts, such as n-butyraldehyde to n-butanol, or benzene tocyclohexane, can be effected with essentially no risk of any part of thecatalyst bed becoming dry. The use of a recycled inert liquid diluent toprevent an overall adiabatic temperature rise over the catalyst bed ofnot more than, typically, about 20° C. to 30° C. in combination with"forced irrigation" of all parts of the catalyst bed by the use of theunconventionally high superficial liquid velocity through the catalystbed prevents the formation of "dry pockets" in the catalyst bed. Theformation of such "dry pockets" where organic vapours and hydrogen arein contact with dry catalyst, in the absence of a continuous liquid flowto remove the heat, can lead to highly exothermic side reactions, e.g.hydrogenolysis of alcohols to hydrocarbons and water, leading to localtemperature runaways, causing poor efficiency of hydrogenation to thedesired product, and reduced catalyst life, as well as reduced catalystutilization efficiency, and even to the formation of tarry materials, orin some cases, to solid coke-like substances.

The hydrogenation zone may comprise an adiabatic reactor, a reactor withan internal cooling coil, or a shell and tube reactor. In the case of ashell and tube reactor the catalyst may be packed in the tubes withcoolant passing through the shell or it may be the shell that is packedwith catalyst with coolant flow through the tubes. The choice of reactordesign will usually be influenced by such factors as the exothermicityof the reaction at the selected inlet concentration of unsaturatedorganic compound or other organic feedstock, the thermal sensitivity ofthe catalyst, and the temperature dependence of any by-product formationreaction, as well as by fluid flow considerations to ensure that evendistribution of gas and liquid within the catalyst volume is obtained.Generally, however, when an adiabatic temperature rise across thecatalyst bed of from about 20° C. to about 30° C. can be accepted, asimple hydrogenation reactor consisting of one or more beds of catalystin a cylindrical vessel with its axis arranged vertically can be usedwith good results. When two or more beds are used in such a reactor thespace between adjacent beds will be largely occupied by the gas phase.The liquid emerging from one bed may with advantage be collected andpassed over a distributor of conventional design before entering thenext bed.

The hydrogen containing gas is generally admixed with the feed solutionupstream from the hydrogenation zone and is partly dissolved therein. Atthe upper end of the hydrogenation zone the concentration of unsaturatedorganic compound or other organic feedstock is at its highest in theliquid phase; hence the rate of hydrogenation is greatest at the upperend of the hydrogenation zone. As the liquid phase passes downwardlythrough the bed of catalyst particles co-currently with the hydrogen itbecomes depleted in respect of hydrogenatable material and to someextent in respect of dissolved hydrogen. The dissolved hydrogen iscontinuously replenished from the gas phase at a rate which is dependentupon the difference between the actual concentration of dissolvedhydrogen and the concentration of dissolved hydrogen at saturation inthe liquid. As a result of the depletion of hydrogen from the gas phasethe partial pressure of any inert gas or gases present rises and thepartial pressure of hydrogen falls as the hydrogen is consumed by thechemical reactions taking place in the hydrogenation zone. Hence at thelower end of the hydrogenation zone the driving force for thehydrogenation reaction can be relatively low. The reaction productexiting the lower end of the hydrogenation zone accordingly usuallystill contains a minor amount of chemically unsaturated or otherhydrogenatable material. Typically the reaction product exiting thehydrogenation zone contains from about 0.01 mole % to about 0.5 mole %,up to about 5 mole % or more of chemically unsaturated or otherhydrogenatable organic material.

As already mentioned, the organic feedstock used as starting materialmay be an unsaturated organic compound that includes two or morehydrogenatable unsaturated groups which may undergo more or lessselective hydrogenation in passage through the hydrogenation zone. Forexample, when an olefinically unsaturated aldehyde (such as2-ethylhex-2-enal) is hydrogenated, the olefinic bond tends to behydrogenated first, before the aldehyde group, so that the saturatedaldehyde (such as 2-ethylhexanal) is a recognisable intermediateproduct. However, some hydrogenation of the aldehyde group may occurprior to hydrogenation of the olefinic linkage, so that2-ethylhex-2-enol is an alternative intermediate product but isgenerally formed in lesser amounts. Each of these intermediates can thenundergo hydrogenation to the desired alcohol product, e.g.2-ethylhexanol.

When an unsaturated organic compound is used as starting material thatcontains only a single hydrogenatable linkage then the unsaturatedhydrogenatable organic material in the reaction product exiting thehydrogenation zone will comprise the unsaturated organic compounditself. However, when an unsaturated organic compound is used asstarting material that contains more than one hydrogenatable unsaturatedlinkage, then the unsaturated hydrogenatable organic material in thereaction product exiting the hydrogenation zone will be selected fromthe starting material and any partially hydrogenated intermediates. Forexample, when hydrogenating 2-ethylhex-2-enal, the hydrogenatableunsaturated organic material in the reaction product may be selectedfrom 2-ethylhex-2-enal, 2-ethylhexanal, 2-ethylhex-2-enol, and a mixtureof two or more thereof.

Generally speaking the depth of the catalyst bed and the hydrogenationconditions in the hydrogenation zone are selected so as to effecthydrogenation of from about 75% to about 99% or more of anyhydrogenatable groups present in the organic feedstock supplied to thehydrogenation zone. Typically the hydrogenation is completed to anextent of from about 85% to about 99.5% in the hydrogenation zone. Inzone cases, however, the extent of hydrogenation in passage through thehydrogenation zone may be higher than this, e.g. 99.8% or more up toabout 99.95%. Such hydrogenation conditions include supply ofhydrogen-containing gas to the upper part of the hydrogenation zone inan amount sufficient to supply an amount of hydrogen that is greaterthan or equal to the stoichiometric quantity required to effect thedesired degree of hydrogenation in the hydrogenation zone. Usually itwill be desirable to limit the supply of hydrogen-containing gas theretoso as to provide as nearly as possible such stoichiometric quantity ofhydrogen and thereby to minimise hydrogen losses in the purge streamfrom the plant. The rate of supply of hydrogen-containing gas to thehydrogenation zone will be mainly dependent upon its composition. Itwill often be preferred to limit the rate of supply so as to provide notmore than about 115% (e.g. up to about 110%), and even more preferablynot more than about 105% (e.g. about 102%), of the stoichiometricquantity required to effect the desired degree of hydrogenation in thehydrogenation zone.

If the hydrogen containing gas is substantially pure hydrogen, e.g. ifit contains about 99.5 mole % or more of hydrogen, then very highdegrees of hydrogenation, exceeding about 99% in suitable cases, can beachieved with the use of a low stoichiometric excess (e.g. about 102%)of hydrogen in a single hydrogenation zone. If, however, the availablehydrogen containing gas is of moderate purity (e.g. one containing about80 to about 90 mole % hydrogen) or of low purity (e.g. one containingless than about 80 mole % hydrogen), then the process can still beoperated using only a low stoichiometric excess of hydrogen by use oftwo hydrogenation zones in series, as taught by WO-A -88/05767 published11th August 1988 the disclosure of which is herein incorporated byreference. Any second or successive hydrogenation zone operating under aco-current flow regime is also desirably operated according to theteachings of the present invention.

The composition of the feed solution will depend upon factors such asthe exothermicity of the hydrogenation reaction, the maximum permissibletemperature rise in the hydrogenation zone, the design of thehydrogenation zone, and the maximum permissible rate of supply to thehydrogenation zone. When operating under adiabatic conditions with anunsaturated organic compound as the organic feedstock, the unsaturatedorganic compound (e.g. aldehyde):inert diluent molar ratio typicallyranges from about 1:3 to about 1:10 and the rate of supply of feedsolution to the hydrogenation zone ranges up to a rate corresponding tosupply of unsaturated organic compound of about 8 moles per liter ofcatalyst per hour or more, e.g. up to about 10 or even 12 moles ofaldehyde or other unsaturated organic compound per liter of catalyst perhour. If, however, provision is made for cooling the hydrogenation zoneas, for example, by use of internal cooling coils within the catalystbed or by use of a shell and tube reactor, then a higher concentrationof unsaturated organic compound can be used; hence in this case theunsaturated organic compound:inert diluent molar ratio typically rangesfrom about 1:1 up to about 1:10.

The inlet temperature to the hydrogenation zone will be at least as highas the threshold temperature for the reaction and will be selected independence on the nature of the hydrogenation reaction. It will normallylie in the range of from about 40° C. to about 350° C., whilst theoperating pressure typically lies in the range of from about 1 bar toabout 300 bar. For example when hydrogenating an aldehyde by the processof the invention the inlet temperature to the hydrogenation zone istypically from about 90° C. to about 220° C. and the pressure istypically from about 5 to about 50 bar.

Besides any remaining hydrogenatable organic feedstock and thehydrogenation product and diluent (if different from the hydrogenationproduct), the liquid reaction product leaving the hydrogenation zonealso contains dissolved inert gases and hydrogen. The gas phase leavingthe hydrogenation zone contains a higher level of inert gases than thehydrogen-containing gas supplied to the upper part of the hydrogenationzone because hydrogen has been removed by the hydrogenation reaction inpassage through the hydrogenation zone.

The reaction product exiting the hydrogenation zone (hereafter sometimescalled "the first-mentioned hydrogenation zone") may be passed through afurther hydrogenation zone in countercurrent to, or in co-current with,a flow of hydrogen-containing gas, in accordance with the teachings ofWO-A-87/07598 published 17th December 1987 or of WO-A-88/05767 published11th August 1988, the disclosure of each of which is herein incorporatedby reference, for the purpose of removing final traces of hydrogenatableorganic material. When any further hydrogenation zone is operated withco-current flow of hydrogen and liquid, it is preferred to operate suchfurther hydrogenation zone also according to the teachings of thepresent invention.

When counter-current flow is used in the further hydrogenation zone, astaught by WO-A-87/07598 published 17th December 1987, the liquid phasefrom the bottom of the first-mentioned hydrogenation zone is fed inliquid form in countercurrent to an upward flow of hydrogen-containinggas. The gas fed to the further hydrogenation zone may have the samecomposition as that supplied to the first-mentioned hydrogenation zone.It is fed to the further hydrogenation zone generally in lesser amountsthan the amount of hydrogen-containing gas supplied to thefirst-mentioned hydrogenation zone. Generally speaking, it should be fedto the further hydrogenation zone in an amount sufficient to provide anat least stoichiometric amount of hydrogen corresponding to the amountof hydrogenatable material remaining in the liquid phase recovered fromthe bottom of the first-mentioned hydrogenation zone. Usually it will bepreferred to supply hydrogen-containing gas to the further hydrogenationzone at a rate sufficient to supply not more than about 115% (e.g. up toabout 110%), preferably not more than about 105% (e.g. about 102%), ofthe stoichiometric quantity of hydrogen required to complete thehydrogenation of the hydrogenatable organic material in the liquid phasefrom the first-mentioned hydrogenation zone.

If desired, the gas fed to the further hydrogenation zone incountercurrent to the liquid flow may be richer in hydrogen than thatfed to the first-mentioned hydrogenation zone. Hence the gas fed to thefirst-mentioned hydrogenation zone may be, for example, a 3:1 molar H₂:N₂ mixture obtained by conventional methods from synthesis gas, whilstthe hydrogen stream to the further hydrogenation zone is a substantiallypure H₂ stream formed by subjecting the same H₂ :N₂ mixture topurification e.g. by pressure swing absorption.

In the further hydrogenation zone the highest H₂ partial pressure existsat the lower end thereof under a counter-current flow regime. Hence thedriving f towards the desired hydrogenation product is maximised in thefurther hydrogenation zone and essentially all of the remainingunsaturated material in the liquid phase exiting the first-mentionedhydrogenation zone is hydrogenated in passage through the furtherhydrogenation zone.

An effluent stream comprising inert gases and hydrogen may be taken fromthe plant between the first-mentioned and further hydrogenation zones inthis preferred process which utilises a counter-current flow regime inthe further hydrogenation zone. This may be passed through a condenserin order to substantially recover any vaporised organic compoundstherein. The resulting condensate is conveniently returned to the top ofthe further hydrogenation zone.

The catalyst beds of the first-mentioned and further hydrogenation zoneswill usually each be supported on a suitable grid. When both beds aremounted in the same vessel, liquid intermediate reaction product fromthe first-mentioned hydrogenation zone may simply be allowed to dropstraight on top of the catalyst bed of the further hydrogenation zonewhen counter-current flow is used in the further hydrogenation zone.Usually, however, it will be desirable to collect and then toredistribute the liquid phase from the first-mentioned hydrogenationzone evenly over the upper surface of the catalyst bed of the furtherhydrogenation zone with the aid of a suitable liquid distributiondevice. In some cases it may be desirable to collect and redistributeliquid within the first-mentioned and/or further hydrogenation zones.

In a preferred process according to the invention for hydrogenation ofan aldehyde the entry temperature to the first-mentioned hydrogenationzone lies in the range of from about 90° C. to about 220° C. and thepressure lies in the range of from about 5 bar to about 50 bar.

In operation of the process of the invention, under steady stateconditions, the composition of the gas (whether dissolved in the liquidphase or present in the gaseous state) exhibits a significant variationbetween different parts of the plant. Thus, for example, the partialpressure of hydrogen is highest in the, or in each, hydrogenation zoneat the respective gas inlet end thereof and lowest at the exit end forgaseous effluent therefrom, whilst the combined partial pressures of anyinert materials present is lowest at the respective gas inlet end tothe, or to each, hydrogenation zone and highest at the exit end forgaseous effluent therefrom. It is thus possible to discharge from thehydrogenation zone a purge gas containing about 50 mole % or more,typically at least about 75 mole %, of inert gases and less than about50 mole % of hydrogen, typically less than about 25 mole % of hydrogen.Under suitable operating conditions it is possible to operate theprocess of the invention so that the effluent gases contain a relativelysmall concentration of hydrogen (e.g. 25 mole % or less) and consistpredominantly of inert gases (e.g. N₂, Ar, CH₄ etc). In this case theeffluent gas stream or streams from the plant is or are relatively smalland consequently hydrogen losses are minimal. In general the compositionand rate of withdrawal of the purge gas stream or streams will bedependent in large part upon the level of inert gases in the hydrogencontaining gas. In the limit, when operating with very pure hydrogen,the solubility of inert gases in the reactor effluent is sufficient topurge such inert gases from the plant and it becomes unnecessary topurge an effluent gas stream from the hydrogenation zone, the inertgases being purged in the course of work up of the hydrogenationproduct.

Because any inert gases present are automatically concentrated in anygaseous effluent stream or streams, it is not necessary on economicgrounds to recycle the gaseous effluents from the hydrogenation zone orzones so as to obtain efficient usage of hydrogen. Recycle of gas isnecessary in conventional co-current or counter-current hydrogenationprocesses in order to achieve efficiency of operation. Moreover, as itis not necessary to recycle a gas stream which contains appreciableconcentrations of inert gases so as to achieve satisfactory economy ofhydrogen consumption, the total operating pressure of the plant cantherefore be reduced although the hydrogen partial pressure ismaintained; hence the construction costs can be reduced as the plant notonly operates at a lower pressure but also no gas recycle compressor isneeded. The absence of a gas recycle compressor, which is in itself anexpensive item of equipment, means also that the civil engineering workassociated with its installation, such as provision of a mounting and acompressor house therefor, is obviated. In addition the ancillary itemsof equipment normally needed when a gas recycle compressor is installed,such as a drive motor, power transformer, and instrumentation, are notrequired. There is also a saving in pipework for the plant as noprovision for recycle of gas is needed. Although it is difficult togeneralise, preliminary calculations suggest that the overall capitalsavings that can be achieved by adopting the process of the inventionfor an aldehyde hydrogenation plant with a throughput of 50,000 tonnesper year can be as much as about 20% compared with the cost of aconventionally designed aldehyde hydrogenation plant. Hence all of thesefactors have a significant effect on both capital and operating costs,both of which are lower for a plant constructed to operate the processof the invention than for conventional co-current or counter-currenthydrogenation plants. Moreover, particularly in the case when a furtherhydrogenation zone is included in the plant as a "polishing" reactor forremoval of the usually small amounts of hydrogenatable organic materialspresent in the liquid phase from the first-mentioned hydrogenation zone,which acts as a "bulk" hydrogenator for hydrogenation of the majority ofthe unsaturated organic compound, the downstream processing of thehydrogenation product is greatly facilitated as the product from theplant is essentially pure hydrogenation product. This also has aprofound and beneficial effect on the capital cost and running costs ofthe product purification section.

In order that the invention may be clearly understood and readilycarried into effect five preferred processes in accordance therewithwill now be described, by way of example only, with reference to FIGS. 1to 5 of the accompanying drawings, each of which is a simplified flowdiagram of a hydrogenation plant constructed in accordance with theinvention, while

FIG. 6 illustrates an experimental hydrogenation apparatus,

FIGS. 7 and 8 plot data obtained from its use,

FIGS. 9 and 10 illustrate a hydrodynamic test rig used to demonstratethe principles underlying the invention, and

FIGS. 11 to 13 summarise data obtained from the rig of FIGS. 9 and 10.

It will be understood by those skilled in the art that FIGS. 1 to 5 arediagrammatic and that further items of equipment such as temperature andpressure sensors, pressure relief valves, control valves, levelcontrollers and the like would additionally be required in a commercialplant. The provision of such ancillary items of equipment forms no partof the present invention and would be in accordance with conventionalchemical engineering practice. Moreover it is not intended that thescope of the invention should be limited in any way by the precisemethods of cooling and heating the various process streams, or by thearrangement of coolers, heaters, and heat exchangers, illustrated inFIGS. 1 to 5. Any other suitable arrangement of equipment fulfilling therequirements of the invention may be used in place of the illustratedequipment in accordance with normal chemical engineering techniques.

Referring to FIG. 1 of the drawings, a stainless steel reactor 1 isprovided with an upper stainless steel grid 2 which supports an upperbed 3 of a granular aldehyde hydrogenation catalyst. This catalyst is aprereduced nickel on alumina catalyst in the form of 1/16 inch (1.6 mm)spheres containing 61% of nickel (calculated as metal) in the 50%reduced form and having a surface area of 140 m² /g as measured by theso-called BET method.

Reactor 1 is of enlarged diameter at its lower end. This enlargeddiameter lower end is fitted with a lower stainless steel grid 4 whichsupports a lower bed 5 of the same nickel catalyst. Thermocouples (notshown) are buried in catalyst beds 3 and 5 and reactor 1 is thermallyinsulated. Steam heating coils (not shown) are provided under thethermal insulation in order to assist in heating reactor 1 at start up.

Layers of open-celled honeycomb grid material (not shown) may be laidone on top of one another on top of grids 2 and 4 as the respective bedis loaded up with catalyst, each layer being offset from the layer belowit so as to assist in even distribution of liquid over the entire bedand to avoid "channelling" of gas through the bed.

The space 6 below lower grid 4 is used to collect liquid emerging fromthe bottom of second bed 5. Such liquid is withdrawn by way of line 7and is recycled by means of pump 8 and lines 9 and 10 through heatexchanger 11 and then through line 12 to a static liquid distributor 13positioned above upper bed 3 at the top of reactor 1.

Reference numeral 14 indicates a feed line for heat exchanger 11 forsupply of a heating medium (e.g. steam) or cooling water as need arises.Heat exchanger 11 can be bypassed by means of by pass line 15, flowthrough which is controlled by means of a valve 16 coupled to atemperature controller 17 which monitors the temperature in line 12.Aldehyde to be hydrogenated is supplied in line 18 and admixed with theliquid exiting heat exchanger 11. The resulting feed solution whichtypically contains about 10% w/w aldehyde is passed by way of line 12 tothe top of catalyst bed 3 at a flow rate corresponding to a superficialliquid velocity down through the catalyst bed 3 of from about 1.5 cm/secto about 3 cm/sec. A liquid intermediate reaction product containingtypically less than about 1,000 ppm aldehyde emerges from the bottom ofbed 3 at substantially the same rate as the flow rate in line 12 andpasses down through catalyst bed 5. Because catalyst bed 5 is of largerdiameter than bed 3 the superficial liquid velocity through bed 5 isless than that through bed 3, typically from about 0.25 cm/sec to about1.0 cm/sec. Alcohol hydrogenation product is withdrawn by way of line 19under the control of valve 20 which is itself controlled by means of alevel controller 21 arranged to monitor the liquid level in bottom space6 of reactor 1.

Hydrogen-containing gas from a pressure swing adsorption unit (notshown) is supplied to reactor 1 in line 22. A major part of the gasflows in line 23 to the top of reactor 1 under the control of a flowcontroller 24 whilst the remainder is fed by way of line 25 under thecontrol of a further flow controller 26 to an upper part of the bottomspace 6 at a point above the liquid level in bottom space 6. Flowcontrollers 24 and 26 are set so that the gas flow rate downwardsthrough catalyst bed 3 at its upper face corresponds to a flow ofhydrogen that is about 105% of the stoichiometric quantity of hydrogenrequired to hydrogenate to alcohol all the aldehyde supplied in line 18.Typically this corresponds to a superficial gas velocity at the uppersurface of bed 3 in the range of from about 1 cm/sec to about 4 cm/sec.A minor amount only of gas flows in line 25, typically ranging fromabout 1% to about 5% of the flow rate in line 23.

A gas purge stream is taken from the space 27 between the two catalystbeds 3 and 5 in line 28. This is passed through a condenser 29 suppliedwith cooling water in line 30. Condensate is collected in drum 31 and isreturned to reactor 1 in line 32. The resulting purge gas stream istaken in line 33 and passed through a further condenser 34 which issupplied with refrigerant in line 35. Pressure control valve 36 is usedto control the pressure within the apparatus and hence the rate ofwithdrawal of purge gas in line 37.

Reference numeral 38 indicates a static liquid distributor fordistributing evenly across the top of lower bed 5 liquid that exitsupper bed 3. Line 39 and valve 40 are used for initial charging of thereactor 1 with liquid.

Reference numeral 41 indicates an optional internal cooling coil whichis supplied with cooling water in line 42.

The use of honeycomb grid material in bed 5 which has been mentionedabove is desirable as an upward flow of hydrogen containing gas iscontacting a downflowing liquid; in this case there is a distincttendency, in the absence of such honeycomb grid material, for the gas toflow up the central axis of the bed and for the liquid to flow down thewalls. The use of honeycomb grid material or of a similar liquid flowdistribution material within catalyst bed 5 helps to obviate thistendency and to promote proper countercurrent flow through bed 5.

The plant of FIG. 2 is generally similar to that of FIG. 1 and likereference numerals have been used therein to indicate like features.

Instead of a single reactor vessel 1 the plant of FIG. 2 has twoseparate reactors 43, 44 each containing a respective catalyst bed 3, 5.Reactor 44 is of larger diameter than reactor 43. Liquid intermediatereaction product emerging from the bottom of first catalyst bed 3collects in the bottom of reactor 43 and passes by way of line 45 to thetop of reactor 44. Purge gas is taken from reactor 43 in line 46 andfrom reactor 44 in line 47 which joins line 46 to form line 48 whichleads in turn to condenser 29. Condensate is returned via line 32 fromdrum 31 to the top of reactor 44.

The apparatus of FIG. 2 permits operation of the two reactors 43 and 44at different pressures; in this case a valve (not shown) can be providedin one or both of lines 46 and 47 and a pump (not shown) can beprovided, if necessary, in line 32.

Referring to FIG. 3 of the drawings, a first reactor 51 is provided withan upper grid 52 which supports an upper bed 53 of a granular aldehydehydrogenation catalyst. This catalyst is a prereduced nickel on aluminacatalyst in the form of 1/16 inch (1.6 mm) spheres containing 61% ofnickel (calculated as metal) in the 50% reduced form and having asurface area of 140 m² /g as measured by the co-called BET method.

First reactor 51 is also fitted with a lower grid 54 which supports alower bed 55 of the same nickel catalyst. Thermocouples (not shown) areburied in catalyst beds 53 and 55 and reactor 51 is thermally insulated.Steam heating coils (not shown) are provided under the thermalinsulation in order to assist in heating reactor 51 at start up.

As in the case of the plant of FIG. 1, layers of honeycomb grid materialcan optionally be introduced into each bed of catalyst as beds 53 and 55are loaded into the reactor 51 in order to assist in promoting evendistribution of liquid throughout the respective bed in operation of theplant.

The space 56 below lower grid 54 is used to collect liquid emerging fromthe bottom of second bed 55. Such liquid is withdrawn by way of line 57and is recycled by means of pump 58 and line 59 through heat exchanger60. It is then fed through line 61 to a second heat exchanger 62 fromwhich it is fed by way of lines 63, 64 to a static liquid distributor 65positioned above upper bed 53 at the top of first reactor 51.

Reference numeral 66 indicates a feed line for heat exchanger 11 forsupply of a heating medium (e.g. steam) or cooling water as need arises.Heat exchanger 62 is provided with a steam heating line 67. Aldehyde tobe hydrogenated is supplied in line 68 and admixed with the liquidexiting heat exchanger 62. This is mainly product alcohol, but stillcontains a minor amount of hydrogenatable material. It acts as a diluentfor the aldehyde. The rate of recycle in line 64 is selected so as toproduce, upon admixture with the incoming aldehyde in line 68, asolution of aldehyde in the product alcohol which typically lies in therange of from about 5 mole % up to about 30 mole % and is selected suchthat the superficial liquid velocity down through catalyst beds 53 and55 is in the range of from about 1.5 cm/sec to about 3 cm/sec.

Part of the recycle stream in line 63 is withdrawn by way of line 69 andis passed by way of line 70 to a static liquid distributor 71 fittednear the top of a second reactor 72.

Hydrogen-containing gas is supplied to first reactor 51 in line 73. Thesource of such hydrogen-containing gas will be described further below.

A gas purge stream is taken from the space 56 below catalyst bed 55 inline 74. This is passed through a condenser 75 supplied with coolingwater in line 76. Condensate is collected in gas-liquid separator 77 andis returned to line 57 in line 78. Reference numeral 79 indicates a misteliminator pad. The resulting purge gas stream is taken in line 80 andis passed through a vent valve 81 which is used to control the pressurewithin the apparatus and hence the rate of discharge of purge gas inline 82.

Second reactor 72 is provided with an upper grid 83 which supports anupper bed 84 of hydrogenation catalyst and with a lower grid 85 whichsupports a lower bed 86 of the same catalyst. The catalyst of beds 84and 86 may be the same as that of beds 53 and 55. Layers of honeycombgrid material may optionally be included in beds 84 and 86 to assist inobtaining even liquid distribution therethrough.

Make up hydrogen-containing feed gas is supplied to the plant in line 87from a pressure swing adsorption unit (not shown), is compressed (ifnecessary) by means of gas compressor 88 and is then passed by way ofheat exchanger 89 and line 90 to the upper end of second reactor 72.Reference numeral 91 indicates a steam heating line. The gas from line90 and the feed in line 70 flow in co-current downwardly through secondreactor 72. The rate of supply of make up gas is controlled so as tocorrespond to about 105% of the stoichiometric quantity of hydrogenrequired to hydrogenate to product alcohol all of the aldehyde suppliedin line 68 after allowance is made for dissolved hydrogen leaving thesystem in the product stream in line 96. This generally corresponds to asuperficial velocity of gas entering the top of catalyst bed 84 in therange of from about 1 cm/sec to about 4 cm/sec. As the feed solutionsupplied in line 70 to second reactor 72 contains only traces ofhydrogenatable organic material, very little hydrogen reacts in passagethrough beds 84 and 86. Substantially all of any hydrogenatable materialremaining in the liquid in line 69 is hydrogenated in passage throughsecond reactor 72. Hence what collects in the space 93 at the bottom ofsecond reactor 72 below catalyst bed 86 is a mixture ofhydrogen-containing gas and product alcohol. This is led in line 94 to aproduct recovery drum 95; hydrogen-containing gas therefrom is led byway of line 73 to the upper end of first reactor 51, as explainedhereinabove. The gas flows into the top of catalyst bed 53 at asuperficial velocity of from about 1 cm/sec to about 4 cm/sec. Liquidproduct alcohol which collects in drum 95 is recovered in line 96 andpassed on for product purification in conventional manner, e.g.distillation in one or more fractional distillation stages.

Second reactor 72 can be operated, as described above, on a once-throughbasis as a single pass reactor. Alternatively the incoming intermediatereaction product in line 69 can be admixed with recycled product fromproduct recovery drum 95. To this end a bypass line 97 is provided toenable recycle to be effected by means of recycle pump 98. This pumpscrude liquid alcohol product by way of line 99 through heat exchanger150 and then via line 151 to a further heat exchanger 152 for recycle inline 153 and admixture with intermediate reaction product in line 69.Reference numerals 154 and 155 indicate heating or cooling lines forheat exchangers 150 and 152 respectively, by means of which temperaturecontrol of the liquid supplied in line 70 can be controlled.

Pump 98 and heat exchangers 150 and 152 can be used at start up of theplant to warm up the catalyst beds 84 and 86 by circulating alcoholthrough reactor 72 prior to introduction of aldehyde to the plant. Heatexchangers 60 and 62 and pump 58 can be used in a similar way tocirculate alcohol through reactor 51 and warm its catalyst beds 53 and55 to the desired starting temperature.

Product alcohol can be supplied to reactor 51 from product recovery drum95, using pump 98, by way of line 156 under the control of valve 157.

If desired, a secondary feed of aldehyde can be supplied by way of line158, e.g. at start up of the plant.

The apparatus of FIG. 3 permits operation of the reactor 51 at adifferent lower pressure than reactor 72; in this case a pressure letdown valve (not shown) can be provided in line 73 and a pump (not shown)can be provided in line 69. Alternatively reactor 72 can be operated ata lower pressure than reactor 51; in this case a compressor (not shown)is provided in line 73 and a valve (also not shown) in line 69.

Instead of two reactor vessels 51 and 72 the plant of FIG. 4 has asingle reactor 101 containing two hydrogenation catalyst beds 102 and103. As with the plant of FIG. 3 each bed may optionally include layersof honeycomb grid material to assist in promoting even distribution ofliquid throughout the bed and to avoid "channelling" of gas through thebed. Catalyst bed 102 constitutes a first hydrogenation zone andcatalyst bed 103 a second hydrogenation zone. Aldehyde to behydrogenated is supplied in line 104 and hydrogen-containing feed gas issupplied from a pressure swing adsorption unit (not shown) in line 105in an amount corresponding to about 105% of the stoichiometric quantityof hydrogen required to hydrogenate all of the aldehyde supplied in line104 to product alcohol.

The aldehyde feed flows from line 104 in line 106 and is admixed with arecycled alcohol stream in line 107. The admixed stream, containingtypically from about 5 mole % to about 30 mole % aldehyde in apredominantly alcohol diluent, is fed in line 108 to a static liquiddistributor 109 above catalyst bed 102. The flow rate is sufficient tocorrespond to a superficial liquid velocity down catalyst bed 102 offrom about 1.5 cm/sec to about 3 cm/sec. Intermediate reaction productis collected at the bottom of reactor 101 and is pumped by way of line110, pump 111 and line 112 to a heat exchanger 113. Then the liquidintermediate reaction product, which contains typically from about 0.1mole % to about 5 mole % chemically unsaturated hydrogenatable organicmaterial, is fed in line 114 to a further heat exchanger 115. Referencenumeral 116 and 117 indicate respective heating or cooling lines forheat exchangers 113 and 115. The liquid intermediate reaction product inline 118 is fed in part in line 107 as the recycle stream to catalystbed 102 and in part via lines 119 and 120 to a further static liquiddistributor 121 fitted above catalyst bed 103. Again, the superficialliquid velocity of the liquid flowing into catalyst bed 103 is fromabout 1.5 cm/sec to about 3 cm/sec.

The chemically unsaturated hydrogenatable organic material remaining inthe intermediate reaction product is substantially all hydrogenated toproduct alcohol in passage through catalyst bed 103. Substantially purealcohol is recovered in line 122 from chimney tray 123 and is pumped bymeans of pump 124 and lines 125 and 126 to a conventional alcoholpurification section (not shown). If desired, part of the productalcohol can be passed by way of line 127 through heat exchangers 128 and129, whose heating or cooling lines are indicated at 130 and 131respectively, to line 132 for recycle to liquid distributor 121.

The hydrogen-containing feed gas in line 105 is compressed as necessaryby means of gas compressor 133, heated in heat exchanger 134, whosesteam heating line is indicated at 135, and supplied in line 136 to thetop of reactor 101 above catalyst bed 103 at a rate corresponding to asuperficial gas velocity at the upper surface of catalyst bed 103 offrom about 1 cm/sec to about 4 cm/sec. Gas emerging from the bottom ofcatalyst bed 103 passes through an orifice 137 in chimney tray 123 andinto catalyst bed 102. As very little hydrogen is consumed in passagethrough bed 103 the superficial gas velocity at the upper surface ofcatalyst bed 102 is similarly in the range of from about 1 cm/sec toabout 4 cm/sec. A purge gas stream is taken from the bottom of reactor101 below catalyst bed 102 in line 138 and is passed through a condenser139 which is supplied with cooling water in line 140. The cooled gas ispassed in line 141 to a gas-liquid separator 142 which is fitted with aspray eliminator pad 143. The purge gas passes out in line 144 throughcontrol valve 145 to a vent line 146. The condensate is returned fromgas-liquid separator 142 to reactor 101 in line 147. Reference numerals148 and 149 represent a bypass line and bypass valve respectively foruse at start up of the plant.

Typical operating conditions in the plants of FIGS. 1 to 4 include useof an inlet temperature to each catalyst bed in the range of from about100° C. to about 130° C. and a pressure of from about 5 bar to about 50bar. In each case the concentration of aldehyde in the feed solution toeach catalyst bed is such as to produce an adiabatic temperature riseacross each bed of no more than about 20° C.

FIG. 5 illustrates a modified form of plant in which an added diluent isused. This form of plant is useful, for example, in the case in whichthe presence of an added adjuvant is desirable, such as ammonia in thehydrogenation of a nitro compound (e.g. nitrobenzene).

Material to be hydrogenated, such as nitrobenzene, is supplied in line201 to a mixing device 202 to which is also fed in line 203 a mixture ofmake up diluent and adjuvant, such as a solution of ammonia in ethanol(containing some water), from line 204 as well as recycleddiluent/adjuvant mixture in line 205. The resulting dilute nitrobenzenesolution is fed to heater 206 in line 207 and admixed with make uphydrogen in line 208. Reference numeral 209 indicates a steam heatingline for heater 206. The mixture of hydrogen, nitrobenzene, ammonia andethanol flows in line 210 to hydrogenation zone 211. This can be asingle reactor or a pair of reactors as used in the plant of one ofFIGS. 1 to 4. As with the plants of FIGS. 1 to 4 layers of open-celledhoneycomb material can be incorporated into the, or into each, catalystbed of hydrogenation zone 211 in order to promote even co-current flowof liquid and gas downward through the bed. The liquid flow rate in line207 is controlled so as to provide a superficial liquid velocity downthrough the or each bed of catalyst of from about 1.5 cm/sec to about 3cm/sec, whilst the gas flow rate in line 208 is adjusted to provide atthe operating pressure and temperature of the plant an amount ofhydrogen equivalent to 115% of the stoichiometrically required amount. Amixture of a hydrogen-depleted purge gas and of an ethanolic anilinesolution, which contains ammonia and water produced by the hydrogenationreaction, is recovered from the bottom of hydrogenation zone 211 in line212. This is fed to a gas liquid separator 213. Gas is purged from theplant in line 214 under the control of valve 215. A cooler 216 issupplied with cooling water in line 217 in order to trap volatilematerials. The liquid phase is led in line 218 to a distillation column219 from which a mixture of ammonia, water and ethanol is recoveredoverhead in line 220 and is condensed by means of condenser 221. Theresulting condensate collects in drum 222; part is returned to column219 in line 223 as a reflux stream whilst the rest is recycled in line224 by means of pump 225 to form the recycle stream in line 205.Reference numeral 226 indicates a gas vent line to condensate drum 222,whilst reference numeral 227 indicates the cooling water supply line forcondenser 221. The bottom product from column 219 in line 228 consistsof substantially nitrobenzene-free aniline containing a minor amount ofethanol and water produced in the reaction. Part is recycled to column219 by way of line 229 and column reboiler 230 whose steam supply lineis indicated at 231. The remainder is passed on for further purificationand storage in line 232.

In a variant of the plant of FIG. 5 mixing device 202 is omitted andlines 201 and 204 are connected to line 224 upstream from pump 225 whichthen serves as a mixing device.

The invention is further illustrated with reference to the followingExamples. Examples 7 and 9 are Comparative Examples and do notillustrate the invention.

EXAMPLES 1 TO 11

The hydrogenation of a C₁₃ aldehyde stream containing 69.98 wt %n-tridecanal, 5.70 wt % 2-methyldodecanal, 0.30 wt % of heavyby-products resulting from aldehyde self condensation reactions and thebalance C₁₂ aliphatic hydrocarbons, was studied in the apparatusdepicted in FIG. 6. This included a reactor 301 made of stainless steeltubing, 2.54 cm internal diameter and 91.4 cm in length, arranged withits axis vertical and fitted with an annular jacket 302 through whichhot oil from a thermostatically controlled bath could be circulated.Reactor 301 contained a bed 303 of catalyst supported on a layer 304 of1.6 mm diameter glass beads 2 cm deep which was itself supported on astainless steel mesh grid 305 some 10 cm above the base of reactor 301.The volume of catalyst bed 303 was 52.3 ml and the catalyst was apre-reduced and air stabilised nickel on alumina catalyst containing 61%w/w of nickel (calculated as metal) in the 50% reduced form and having asurface area of 140 m² /g as measured by the so-called BET method. Thephysical form of the catalyst was near spherical granules of a nominal1/16 inch (1.6 mm) diameter; the actual size range limits of theparticles was from 1.4 mm to 2.36 mm as determined by sieve analysis.The upper portion of reactor 301 was filled with a layer 306 of 1.6 mmdiameter glass beads; this layer 306 ensured that the temperature of thefeed solution and entrained hydrogen supplied to catalyst bed 303 couldbe controlled to a preselected value.

Reactor 301 was also fitted with a thermocouple pocket 307 of smalldiameter for a thermocouple 308. During the packing procedure it wasdetermined that the depth of catalyst bed 303 was 10.5 cm. Liquid couldbe withdrawn from the bottom of reactor 301 in line 309 by means of pump310 and recycled to the top of reactor 301 in line 311. The rate ofrecycle of liquid in line 311 could be measured using a mass flow meter(not shown). Aldehyde feed could be supplied to the apparatus from aburette (not shown) in line 312 by means of a feed pump (not shown).Hydrogen could be supplied from a storage cylinder via a pressure letdown valve and a flow controller (neither of which is shown) in line313. A mixture of gas and liquid could also be withdrawn from reactor301 by means of an overflow pipe 314 and passed in line 315 to agas/liquid separation vessel 316. Pressure control valve 317 allowed apurge gas stream to be let down to atmospheric pressure and passed inline 318 to a wet gas meter (not shown) before being vented to theatmosphere. Liquid product could be removed from the system in line 319by means of a pressure let down valve 320 operating under the influenceof a liquid level controller 321. Samples of this liquid product wereanalysed by gas-liquid chromatography from time to time. Such analysiswas repeated after any change in operating conditions had been effecteduntil the results showed that steady state conditions had beenre-established. The whole apparatus was positioned in a fume cupboardsupplied with warm air at 40° C. to eliminate any danger of blockage oflines due to solidification of n-tridecanol (m.p. 32-33° C.).

After purging the apparatus with nitrogen approximately 120 ml of C₁₃alcohol were charged to the apparatus by means of line 312, thecirculating hot oil flow was established at a temperature of 120° C.,and pump 310 was set into operation. This quantity of liquid wassufficient to fill the bottom of reactor 301. A flow of hydrogen wasestablished through the apparatus and then the system was brought up tooperating pressure and the aldehyde feed pump started. The results arelisted in Table 1. All Examples were carried out using circulating oilat 120° C. and in each case, except Example 7 and especially Example 9when thermocouple 308 indicated an incipient temperature runaway, thetemperature of the catalyst bed 303 remained within 5° C. of 120° C. H₂flow rates are measured in "normal" liters per hour (i.e. liters of gasat 0° C. and 1 bar).

                                      TABLE 1                                     __________________________________________________________________________                        Liquid                                                         Aldehyde                                                                           H.sub.2                                                                            H.sub.2 purge                                                                      Recycle   Bed %  -n-aldehyde                                                                        % "heavies"                         Example                                                                            feed rate                                                                          feed rate                                                                          rate rate SLV  Temp.                                                                             in reactor                                                                            in reactor                          No.  (ml/hr)                                                                            (l/hr)                                                                             (l/hr)                                                                             (l/hr)                                                                             (cm/sec)                                                                           (°C.)                                                                      effluent (w/w)                                                                        effluent (w/w)                      __________________________________________________________________________    1    240  38.7 19.8 25.8 1.58 124.5                                                                             10.3    2.98                                2    120  31.4 19.8 25.8 1.57 123.9                                                                              4.26   1.36                                3     60  26.7 19.8 25.8 1.57 123.1                                                                              1.86   1.18                                4     30  23.9 19.8 25.8 1.57 123.0                                                                              0.94   1.08                                5    480  41.9 19.8 25.8 1.59 124.7                                                                             31.2    6.15                                6    480  62.8 39.5 25.8 1.59 125.0                                                                             30.4    6.26                                7    480  42.5 19.8 13.0 0.82 131.2                                                                             31.3    6.76                                8    480  24.9  3.9 25.8 1.59 124.8                                                                             33.6    6.38                                9    480  45.5 19.8  5.1 0.34 143.9                                                                             26.3    6.82                                10    60  27.1 19.8 25.8 1.57 123.3                                                                              1.56   1.44                                __________________________________________________________________________     Note: The term "SLV" means superficial liquid velocity and is calculated      assuming a density of 0.75 g/cc at reactor inlet conditions and 0.83 g/cc     at room temperature for the reactor inlet feed solution                  

As the recycle rate in line 311 is known and the n-aldehydeconcentration, i.e. [--CHO]_(exit), in the liquid being recycled is alsoknown and as the feed rate and aldehyde concentration in the materialsupplied in line 312 are also known, it is readily possible to calculatethe n-aldehyde inlet feed concentration, i.e. [--CHO]_(inlet), for eachExample. From these figures was calculated, in each case, the meann-aldehyde concentration, i.e. [--CHO]_(mean), in the reactor, accordingto the equation: ##EQU1## The mean n-aldehyde concentration is tabulatedin Table 2 against the percentage change in n-aldehyde concentration{Δ[--CHO]} from one end of the reactor to the other. These dataobservations are plotted in FIG. 7.

                  TABLE 2                                                         ______________________________________                                                                        Liquid                                        Example  [--CHO].sub.mean                                                                           Δ [--CHO]                                                                         recycle rate                                  No.      (% w/w)      (% w/w)   (l/hr)                                        ______________________________________                                        1        10.58        0.55      25.8                                          2         4.41        0.30      25.8                                          3         1.94        0.16      25.8                                          4         0.98        0.08      25.8                                          5        31.56        0.71      25.8                                          6        30.76        0.72      25.8                                          7        31.99        1.38      13.0                                          8        33.93        0.66      25.8                                          9        28.18        3.76       5.1                                          10        1.64        0.16      25.8                                          ______________________________________                                    

Examples 1 to 5 and 10 were all carried out with a liquid recycle rateof 25.8 1/hr and a hydrogen purge rate of 19.8 1/hr so that these datadefine the relationship between the amount of n-aldehyde converted inpassage through reactor 301, i.e. Δ[--CHO], and the n-aldehydeconcentration, [--CHO]_(mean), within the reactor 301 under theseconditions of hydrogen flow and liquid recycle rate. A considerablereduction in hydrogen purge flow rate to 3.9 1/hr makes very littledifference to the amount of n-aldehyde converted in passage throughreactor 301, i.e. Δ[--CHO], as can be seen by comparison of Examples 5and 8. A large increase in hydrogen purge flow rate to 39.5 1/hr makesvery little difference to the amount of n-aldehyde converted in passagethrough reactor 301, i.e. Δ[--CHO], as is readily apparent by comparisonof Examples 5 and 6. In contrast a reduction in liquid recycle rate,although increasing the conversion of n-aldehyde in passage throughreactor 301, i.e. Δ[--CHO], as shown by Examples 7 and 9, caused amarked increase in catalyst bed temperature, as detected by thermocouple308, despite the use of circulating oil at 120° C. in jacket 302. Thisincipient temperature runaway was further accompanied by an increase in"heavies" formation.

The data defining the curve of FIG. 7 represent a scan of differenthorizontal segments of catalyst in a large reactor and can be used tocalculate the depth of catalyst bed required for a commercial rectoroperating under appropriate conditions including aldehyde concentration,flow rate and temperature according to the teachings of the invention.

Comparison of the relative amounts of aldehyde converted over thereactor system calculated from the flow rates and aldehyde concentrationchanges across the reactor in Examples 7 and 9, using Example 5 as areference, shows that virtually the same amount of aldehyde is convertedin the reactor system in Example 7 as in Example 5, despite asignificant increase in catalyst temperature and some increase in heavyby-products production. Comparison of Examples 5 and 9 show that anincrease of only about 12% in aldehyde conversion by the reactor systemhas been gained at the expense of an unacceptable temperature rise andincrease in heavy by-products formation. Example 9 in some measurerepresents the situation arising in a "local low flow volume element" ofa large catalyst bed operated at low superficial liquid velocities.These comparisons illustrate that the space time productivity of thecatalyst is maintained at high liquid superficial velocities and thatpotentially dangerous temperature excursions with consequent loss ofcatalyst activity and selectivity are obviated using the process of theinvention.

EXAMPLES 11 TO 36

The apparatus of FIG. 6 was charged with 58 ml of the same catalyst andwas used to investigate further the hydrogenation of the same C₁₃aldehyde feedstock that was used in Examples 1 to 11. The reactionconditions and the results obtained are summarised in Table 3. InExamples 34 to 36 the C₁₃ aldehyde feedstock was diluted withn-tetradecane. In each case the liquid recycle rate was maintained at 281/hr, thus ensuring that the superficial linear liquid velocity throughthe reactor was at least 1.5 cm/sec.

FIG. 8 summarises the results of Examples 31 to 36. This is a graph ofthe amount of aldehyde converted per hour in the apparatus plottedagainst the concentration of aldehyde in the liquid phase exiting thereactor. The numerals on the graph indicate the numbers of therespective Examples. It will be seen that two separate curves can beplotted, one representing the data obtained when no diluent (i.e.n-tetradecane) has been added and the other when a diluent is used.

                                      TABLE 3                                     __________________________________________________________________________         Reactor                                                                            Aldehyde                   %  -n-aldehyde in                                                                      % "heavies"                     Example                                                                            Pressure                                                                           Feed Rate                                                                            -n-tetradecane                                                                      H.sub.2 Exit Flow                                                                    Reactor Bed                                                                          reactor effluent                                                                       in reactor                      No.  (bar)                                                                              (l/hr)                                                                              added (l/hr)                                                                         (l/hr) Temp. (°C.)                                                                   (w/w)    effluent (w/w)                  __________________________________________________________________________    11   18.24                                                                              0.06  --      63    122.5  1.93     3.22                            12   18.24                                                                              0.24  --     254    123.4  11.58    5.79                            13   25.13                                                                              0.24  --     254    123.4  10.34    5.13                            14   25.13                                                                              0.12  --     109    122.6  4.69     4.34                            15   25.13                                                                              0.12  --      86    133.2  3.14     3.70                            16   25.13                                                                              0.24  --     176    133.7  7.65     5.01                            17   18.24                                                                              0.24  --     193    134.1  8.64     6.56                            18   18.24                                                                              0.12  --     100    133.3  3.75     5.16                            19   21.68                                                                              0.18  --     150    128.8  6.96     5.44                            20   21.68                                                                              0.12  --     107    128.4  4.26     4.82                            21   21.68                                                                              0.24  --     223    129.0  10.07    6.16                            22   25.13                                                                              0.18  --     140    133.6  5.86     5.41                            23   25.13                                                                              0.18  --     183    123.1  8.52     5.79                            24   18.24                                                                              0.18  --     206    123.2  9.37     7.08                            25   18.24                                                                              0.18  --     160    133.5  6.91     6.74                            26   18.24                                                                              0.06  --      57    132.6  1.80     5.07                            27   25.13                                                                              0.06  --      57    132.6  1.58     4.26                            28   18.24                                                                              0.12  --     140    124.1  5.93     6.35                            29   18.24                                                                              0.06  --      73    123.9  2.51     5.22                            30   18.24                                                                              0.24  --     280    124.6  13.18    8.00                            31   18.24                                                                              0.06  --      73    124    2.65     5.22                            32   18.24                                                                              0.12  --     140    124    6.07     6.35                            33   18.24                                                                              0.24  --     280    124    13.32    8.00                            34   18.24                                                                              0.06  0.06    57    124    1.75     2.75                            35   18.24                                                                              0.12  0.12    90    124    4.08     3.21                            36   18.24                                                                              0.24  0.24   177    124    8.43     3.90                            __________________________________________________________________________

Regression analysis of the rate of conversion (R_(N)) of n-aldehyde toproducts (expressed as gm moles of C₁₃ aldehyde converted/liter ofcatalyst/hr) produced an equation of the following form. ##EQU2## whereR_(N) =gm moles of n-aldehyde hydrogenated to products/liter ofcatalyst/hr

T° K.=average catalyst and temperature

RxBar=Reactor pressure (bar)

%NALD=Mean % n-aldehyde in reactor (calculated)

ALH₂ =Calculated actual liters/hr of hydrogen exiting from the bottom ofthe catalyst bed at reactor pressure and temperature

%HVY=% "heavies" in the reactor exit stream

    ______________________________________                                                             Standard Error                                           Coefficient          of Coefficient                                           ______________________________________                                        a = 0.156            0.122                                                    E = -4867.78         255.6                                                    b = 0.837            0.0867                                                   c = 0.0179           0.111                                                    d = 0.4497           0.0356                                                   A is a constant =    345756                                                   e = the base for natural logarithms                                                                (i.e. 2.71828 . . . )                                    ______________________________________                                    

The validity of the above rate equation is shown in Table 4 wherepredicted rates versus actual rates, in gm moles/1 catalyst/hr, arecompared.

                  TABLE 4                                                         ______________________________________                                                               Rate Predicted by Rate                                 Example No.                                                                            Observed Rate Equation                                               ______________________________________                                        11       2.99          2.82                                                   12       10.15         9.77                                                   13       10.33         10.12                                                  14       5.61          5.44                                                   15       5.80          5.86                                                   16       10.71         10.88                                                  17       10.71         10.34                                                  18       5.80          5.56                                                   19       8.25          8.14                                                   20       5.61          5.62                                                   21       10.33         10.58                                                  22       8.25          8.38                                                   23       8.06          8.02                                                   24       7.87          7.62                                                   25       8.25          8.28                                                   26       3.00          2.95                                                   27       3.00          3.01                                                   28       5.61          5.60                                                   29       2.99          2.94                                                   30       9.77          10.06                                                  31       2.92          3.13                                                   32       5.54          5.69                                                   33       9.81          9.98                                                   34       2.88          2.93                                                   35       5.35          5.55                                                   36       9.20          9.40                                                   ______________________________________                                    

This analysis of Examples 11 to 36 shows that:

(a) Hydrogen flow has little or no positive effect on the rate ofhydrogenation under these liquid flow velocity conditions;

(b) Reactor pressure (i.e. hydrogen pressure) has a minor positiveeffect on the reaction rate and is of poor statistical significance(over the pressure range used 18.24 to 25.13 bar); and

(c) "Heavies" are catalyst inhibitors.

These conclusions substantiate in a more rigorous way the insensitivityof the reaction kinetics to the rate of hydrogen passing through thecatalyst bed which can be noted from comparison of Examples 5 and 6 andof Examples 5 and 8. Also the rate equation describes the effect of theprocess conditions on the catalyst in a differential manner; suitableintegration of the equation over the depth of a commercial catalyst bedwill provide a valuable prediction of the bed's performance.

The plants of FIGS. 1 to 5 and the operating techniques described aboveare generally applicable to hydrogenation of organic materials. It willaccordingly be readily apparent to the skilled reader that the teachingsof the invention can be practised with a wide variety of hydrogenationreactions other than the aldehyde hydrogenation reaction specificallydescribed in relation to FIGS. 1 to 4 of the accompanying drawings andthe nitrobenzene hydrogenation reaction described in relation to FIG. 5of the drawings.

EXAMPLE 37

Examples 1 to 36 used experimental systems where reactors of smalldiameter (2.54 cm) were used. Commercial reactors of much largerdiameter are necessary in order to achieve the necessary productionrates. Therefore the distribution of gas and liquid passing inco-current downflow through a much larger bed of particulate solid wasinvestigated in an apparatus which is illustrated in FIGS. 9 and 10.This comprises a rectangular section column 401 which was constructedfrom 1.25 mm thick "Perspex" (Registered Trade Mark) sheet so as toenable its contents to be viewed. Partitions 402 near its base dividedthe base of the column 401 into six bays 403, each of which had acorresponding outlet line 404 for water and an outlet line 405 for air.Reference numeral 406 indicates a perforated support for a bed 407 ofparticles intended to simulate a hydrogenation catalyst. Bed 407consisted of impervious ceramic balls of nominal size 2.4 to b 4 mm,more than 80% of which were 3 mm or less in diameter. Water was suppliedin line 408 to a bar distributor 409 above the top of the bed 407,whilst air was fed in line 410 from a compressor (not shown) to inlets411 at the top of column 401. Bed 407 measured approximately 460 mm×75mm×1,425 mm and was topped with a layer of 12.7 mm diameterpolypropylene balls approximately 200 mm deep which was intended toenhance the uniformity of distribution of the water over the top of bed407. The water that was collected in each bay 403 was conducted along aline 404 of standard length to a corresponding turbine meter in a bank412 of turbine meters, each receiving water from a respective bay 403.Similarly air from each bay 403 was conducted along a line 405 ofstandard length to a corresponding turbine meter in a bank 413 of suchturbine meters, each receiving air from a respective bay 403. Asindicated by reference numerals 414 and 415 the signals from the twobanks of meters 412 and 413 were transmitted to respective data loggers(not shown). By providing lines 404 of essentially identical length anddiameter for water and lines 405 similarly of essentially identicallength and diameter for the air flow from each bay 403 it was ensuredthat, so far as possible, the risk of the air and water flow measurementsystems interfering with the measurements of flow through the bed 407was avoided. However, at low air flow rates, of the order 2 to 3 litersper minute, the air flow measurement turbines of bank 413 becameinaccurate and/or inoperative. Accordingly the corresponding airdistribution measurements have no significance in this low air flowrange. From the meters of bank 412 the water was collected in a tank 414and recirculated to the top of the apparatus by pump 415.

Measurements were made with water flow rates in line 408 of 30 to 55liters per minute and air flow rates in line 410 of 59 to 5 liters perminute. These flow rates were chosen to simulate a range of flow rateslikely to be encountered in a commercial hydrogenation reactor operatedin accordance with the teachings of this invention and correspond to aliquid phase superficial velocity of 1.43 to 2.63 cm/sec and a gas phasesuperficial velocity of 0.096 to 2.01 cm/sec.

The distribution of fluid across the bed 407 was calculated as follows:

For each fluid:

Average flow=sum of flows/6

Variance=[Average flow-measured flow]

Average variance=sum of variances/6

(It should be noted that the variance was always recorded as a positivenumber).

The results are recorded in Tables 5 to 7 and plotted in FIGS. 11 to 13.

At the higher gas and liquid flow rates the operation of a highlydispersed gas/liquid regime was clearly shown.

In those cases where active liquid/air bubble movement was visuallyobservable no static regions of the bed were evident; the phase in agiven bed void was replaced by the other phase at apparently randomintervals.

From the results obtained it would appear that the efficiency of phasedistribution (as measured by the variance from average flow per port) isa function of throughput. That is, higher air/water flows (and hence asteeper pressure gradient) lead to a better gas/liquid distribution.This observed effect is undoubtedly enhanced by the poorer accuracy ofthe measuring devices at low flow rates and also by the increasingeffect of any fortuitous physical variations between the six gas/liquidcollection and separation ports. It is therefore highly probable thatthe actual distribution is always better than the observed distribution.It should also be noted that the corners of the rig of FIGS. 9 and 10provide a low resistance fluid path to the left-hand and right-hand bays403 (as illustrated) for geometrical reasons; this effect will also addto the variance observed. A circular cross section catalyst bed willgive better gas/liquid distributions than those observed with arectangular cross section bed.

These gas/liquid distribution studies show that effective gas and liquidco-current downflow hydrogenation reactions can be achieved withoutusing large excesses of hydrogen containing gases.

                  TABLE 5                                                         ______________________________________                                        Water flow in 30-36 liter/min range                                           Air        Air         Water                                                  l/min      % av. variance                                                                            % av. variance                                         ______________________________________                                         9.8       42.9        6.8                                                    10.9       19.3        3.9                                                    14.8       24.8        6.5                                                    19.7       21.9        5                                                      29.4       15.6        3.6                                                    39.1       14.2        10.4                                                   42.4        4.2        4.4                                                    49.5       12.7        9.3                                                    ______________________________________                                    

                  TABLE 6                                                         ______________________________________                                        Water flow in 44-46 liter/min range                                           Air        Air         Water                                                  l/min      % av. variance                                                                            % av. variance                                         ______________________________________                                         9.3       34.8        5.3                                                    10         36.6        4.2                                                    19.2       20.8        5.9                                                    20.2       16.8        4                                                      28.5       14          6                                                      30         11.7        4.4                                                    37.9       11.8        5.2                                                    40         10.2        3.9                                                    44.9       12.3        4.5                                                    47.7        9.2        4.2                                                    48.4        9.1        3.1                                                    ______________________________________                                    

                  TABLE 7                                                         ______________________________________                                        Water flow in 53-56 liters/min range                                          Air        Air         Water                                                  l/min      % av. variance                                                                            % av. variance                                         ______________________________________                                        11.4       21.1        5.6                                                    13.1       23.4        6.5                                                    20.2       13.5        5.2                                                    33.1       8.9         2.9                                                    33.3       8.4         5.4                                                    43         5.6         3.4                                                    43.3       5.8         3.3                                                    43.7       7.5         4.4                                                    50.3       3.8         3.6                                                    59.8       7.6         3.6                                                    ______________________________________                                    

We claim:
 1. A liquid phase catalytic hydrogenation process in which anorganic feedstock is contact with hydrogen in the presence of a solidhydrogenation catalyst under hydrogenation conditions to produce ahydrogenation product, which process comprises:passing a feed solutionof the organic feedstock in an inert diluent therefor downwardly inco-current with a hydrogen-containing gas through a hydrogenation zonecontaining a bed of a particulate hydrogenation catalyst whose particlessubstantially all lie in the range of from about 0.5 mm to about 5 mm,maintaining the bed of catalyst particles under temperature and pressureconditions conducive to hydrogenation, recovering from a bottom part ofthe bed a liquid phase containing the hydrogenation product, controllingthe rate of supply of the feed solution to the bed so as to maintain asuperficial liquid velocity of the liquid down the bed in the range offrom about 1.5 cm/sec to about 5 cm/sec, and controlling the rate ofsupply of the hydrogen-containing gas to the bed at the chosen rate ofsupply of feed solution so as to set up a pressure drop across the bedof at least about 0.1 kg/cm² per meter of bed depth, so as to maintainat the top surface of the bed of catalyst particles of flow ofhydrogen-containing gas containing from 1.00 to about 1.15 times thestoichiometric quantity of hydrogen theoretically necessary to convertthe organic feedstock completely to the hydrogenation product and so asto ensure that all parts of the bed are subjected to forced irrigationwith liquid containing entrained bubbles of hydrogen-containing gas. 2.A process according to claim 1, in which the hydrogenation conditionsinclude use of a pressure of from about 1 bar to about 300 bar and of atemperature of from about 40° C. to about 350° C.
 3. A process accordingto claim 1 or claim 2, in which the organic feedstock comprises analdehyde containing from 2 to about 20 carbon atoms, and in which thehydrogenation product is an alcohol containing from 2 to about 20 carbonatoms.
 4. A process according to claim 3, in which the hydrogenationconditions include use of a pressure of from about 5 bar to about 50 barand of a temperature of from about 90° C. to about 220° C.
 5. A processaccording to claim 1 or claim 2, in which the organic feedstock is anunsaturated hydrocarbon.
 6. A process according to any one of claims 1,or 2, in which the superficial liquid velocity down the bed is fromabout 1.5 cm/sec up to about 3 cm/sec.
 7. A process according to any oneof claims 1, or 2, in which the hydrogen containing gas contains atleast about 90 mole % of hydrogen.
 8. A process according to any one ofclaims 1, or 2, in which the hydrogenation zone is operated underadiabatic conditions and the concentration of organic feedstock in thefeed solution is selected to produce an adiabatic temperature rise inpassage through the bed of not more than about 30° C.
 9. A processaccording to any one of claims 1, or 2, in which the particles of theparticulate hydrogenation catalyst substantially all lie in the range offrom about 0.5 mm to about 3 mm.
 10. A process according to claim 3, inwhich the superficial liquid velocity down the bed is from about 1.5cm/sec up to about 3 cm/sec.
 11. A process according to claim 5, inwhich the superficial liquid velocity down the bed is from about 1.5cm/sec up to about 3 cm/sec.
 12. A process according to claim 3, inwhich the hydrogen-containing gas contains at least about 90 mole % ofhydrogen.
 13. A process according to claim 5, in which thehydrogen-containing gas contains at least about 90 mole % of hydrogen.14. A process according to claim 6, in which the hydrogen-containing gascontains at least about 90 mole % of hydrogen.
 15. A process accordingto claim 3, in which the hydrogenation zone is operated under adiabaticconditions and the concentration of organic feedstock in the feedsolution is selected to produce an adiabatic temperature rise in passagethrough the bed of not more than about 30° C.
 16. A process according toclaim 5, in which the hydrogenation zone is operated under adiabaticconditions and the concentration of organic feedstock in the feedsolution is selected to produce an adiabatic temperature rise in passagethrough the bed of not more than about 30° C.
 17. A process according toclaim 6, in which the hydrogenation zone is operated under adiabaticconditions and the concentration of organic feedstock in the feedsolution is selected to produce an adiabatic temperature rise in passagethrough the bed of not more than about 30° C.
 18. A process according toclaim 7, in which the hydrogenation zone is operated under adiabaticconditions and the concentration of organic feedstock in the feedsolution is selected to produce an adiabatic temperature rise in passagethrough the bed of not more than about 30° C.
 19. A process according toclaim 3, in which the particles of the particulate hydrogenationcatalyst substantially all lie in the range from about 0.5 mm to about 3mm.
 20. A process according to claim 5, in which the particles of theparticulate hydrogenation catalyst substantially all lie in the rangefrom about 0.5 mm to about 3 mm.
 21. A process according to claim 6, inwhich the particles of the particulate hydrogenation catalystsubstantially all lie in the range from about 0.5 mm to about 3 mm. 22.A process according to claim 7, in which the particles of theparticulate hydrogenation catalyst substantially all lie in the range ofabout 0.5 mm to about 3 mm.
 23. A process according to claim 8, in whichthe particles of the particulate hydrogenation catalyst substantiallyall lie in the range from about 0.5 mm to about 3 mm.